Industrial and hydrocarbon gas liquefaction

ABSTRACT

A method for liquefaction of industrial gases or gas mixtures (hydrocarbon and/or non-hydrocarbon) uses a modified aqua-ammonia absorption refrigeration system (ARP) that is used to chill the gas or gas mixture during the liquefaction process. The gas may be compressed to above its critical point, and the heat of compression energy may be recovered to provide some or all of the thermal energy required to drive the ARP. The method utilizes a Joule Thomson (JT) adiabatic expansion process which results in no requirement for specialty cryogenic rotating equipment. The aqua-ammonia absorption refrigeration system includes a vapour absorber tower (VAT) which permits the recovery of some or all of the heat of solution and heat of condensation energy in the system when anhydrous ammonia vapour is absorbed into a subcooled lean aqua-ammonia solution. The modified ARP with VAT may achieve operating pressures as low as 10 kPa which results in ammonia gas chiller operating temperatures as low as −71 C.

FIELD OF THE INVENTION

The present invention relates to systems and methods for theliquefaction of industrial or hydrocarbon gases or gas mixtures.

BACKGROUND OF THE INVENTION

Industrial gases such as CO₂, H₂S, N₂, O₂, H₂, He, Ar, air and othergases, and hydrocarbon gases such as methane, ethane, propane, ethyleneand other hydrocarbon gases, or mixtures of gases, are traditionallyliquefied utilizing refrigeration cycles based on well-known Carnotrefrigeration or Turbo-Expander cycles. The cryogenic temperaturesachieved during these industrial processes which enable liquefaction canrequire complex cascaded refrigeration cycles that are capital, energy,and operating cost intensive.

Accordingly, there is a need in the art for alternative methods ofliquefying industrial and hydrocarbon gases or gas mixtures that may berelatively energy efficient, economical and practical to implement.

SUMMARY OF THE INVENTION

In one aspect, the invention comprises a method for liquefying a gas,comprising the following non-sequential steps:

-   -   a. receiving a gas having an inlet pressure and compressing or        decompressing the gas to a desired pressure;    -   b. chilling the gas through at least one absorption chiller;    -   c. adiabatically reducing the pressure of the gas to liquefy at        least a portion of the gas;    -   d. heating a rich aqua-ammonia fluid in a rectifier to liberate        ammonia gas using one or a combination of trim heat or heat of        compression recovered from step (a) if the gas is compressed in        step (a), producing a lean aqua-ammonia fluid;    -   e. subcooling the lean aqua-ammonia and circulating to the top        of a vapour absorption tower;    -   f. condensing the ammonia gas from the rectifier and flashing        the liquid ammonia to produce chilled ammonia gas for use in the        at least one absorption chiller;    -   g. absorbing ammonia gas from the at least one absorption        chiller into the lean aqua-ammonia in the vapour absorption        tower to produce the rich aqua-ammonia for step (d).

The gas may comprise an industrial gas or a hydrocarbon gas, or anymixture of industrial or hydrocarbon gases. The method may result in theliquefaction of at least one component of the gas, a portion of the gas,or substantially all of the gas.

In another aspect, the invention may comprise a gas liquefaction systemcomprising a receiving stage for receiving an inlet gas at an inletpressure, a chilling stage comprising an absorption refrigeration loopfor chilling the received gas, and a liquefaction stage comprising a JTvalve for at least partially liquefying the chilled gas. In oneembodiment, the system may further comprise a compression stage forcompressing the gas to the desired pressure, and a heat of compressionenergy recovery stage for transferring heat from the compression stageto the absorption refrigeration loop. In another embodiment, the systemmay comprise a gas recycle stage for recycling non-liquefied componentsof the gas in a low pressure vapour recycle loop, which loop furtherchills the compressed and chilled gas, and which is then directed to thecompression stage.

In one embodiment, the absorption refrigeration loop comprises arectifier and a vapour absorption tower.

BRIEF DESCRIPTION OF THE DRAWINGS

The invention will now be described by way of exemplary embodiments withreference to the accompanying simplified, diagrammatic drawings.

FIG. 1 is a schematic depiction of one embodiment of the presentinvention.

FIG. 2 is a process flow diagram (PFD) of one embodiment, where the gasis compressed to less than the critical pressure for the gas.

FIG. 3 is a Mollier Chart for carbon dioxide (CO₂) utilizing oneembodiment of the present invention. This and other Mollier Charts showspecific Enthalpy-Pressure Charts as provided by ChemicalogicCorporation USA.

FIG. 4 is a process flow diagram (PFD) utilizing gas liquefaction method2 for a sweet natural gas at 170 kPa inlet pressure, water saturatedwith 2% CO₂ and 98% CH₄. The liquefaction cycle uses a single flashliquefaction application to a storage of pressure of 170 kPa.

FIG. 5 is a process flow diagram (PFD) utilizing the modified absorptionrefrigeration cycle noting key equipment and process data points. Theprocess shows key components for 4 stage NH₃ chiller system, vapourabsorber tower (VAT), lean solution chiller, waste heat exchangers,generator, rectifier column, reflux condenser (dephlegmator), ammoniacondenser, and other ancillary equipment

FIG. 6A is a Mollier Chart for methane (CH₄) utilizing a liquefactioncycle showing one embodiment of the present invention. FIG. 6B is aMollier Chart for methane utilizing an alternative embodiment, with anoptional high-pressure feed.

FIG. 7 is a Mollier Chart for anhydrous ammonia (NH₃) which shows thethermodynamic points for a 4 stage liquefier chiller system which showthe pressure and temperature of the anhydrous ammonia vapour as itreturns to the VAT. Ambient cooling system temperature for this exampleassumes a condensing temperature of 22° C.

FIG. 8 is a PTX drawing for aqua-ammonia solution that represents theoperating points, in particular the key process operating pressures,temperatures and solution concentration through the VAT and theremainder of the modified absorption cycle as utilized in the invention.The PTX graph for Aqua-Ammonia was plotted utilizing process data fromPROMAX™ process simulator.

FIG. 9 is a process flow diagram (PFD) of one embodiment, where thefinal gas liquefaction cooling occurs in a liquefied gas vapourizationheat exchanger.

FIG. 10 is a Mollier Chart for the liquefaction of air utilizing oneembodiment of a liquefaction cycle as shown in FIG. 9.

DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS

In physics, the term “gas” includes a state of matter where a substancehas perfect molecular mobility and the property of indefinite expansion.As used herein, a “gas” include substances which are gases at standardtemperature and pressure, such as CO₂, H₂S, N₂, O₂, H₂, He, Ar, air, orhydrocarbon gases such as methane, ethane, propane, ethylene and otherhydrocarbon gases, or any mixture of gases. As used herein, the term“liquefied gas” means any gas or mixture of gases, that has beenliquefied for sale, disposal or use for commercial, research orindustrial purposes.

As used herein, the term “JT valve” or “JT throttling valve” means a gasvalve adapted to allow the adiabatic expansion of gas in accordance withthe Joule-Thompson effect. JT valves are well known in the art, and arecommercially available.

As used herein, the term “low pressure separator” or “LPS” means aseparating vessel that operates at a specified lower pressure andtemperature downstream of a “JT” throttling valve, such that a liquefiedgas can be removed from the flow path or processed further within theflow path.

As used herein, the term “high pressure separator” or “HPS” means aseparating vessel that operates at the desired pressure for gas chillingand is located upstream of the JT throttling valve,

As used herein, the term “dense phase” as it relates to any gas or gasmixtures means the state of a gas resulting from its compression aboveits cricodenbar, which is the maximum pressure above which the gascannot be formed into the gas phase, regardless of temperature, at atemperature within a range defined by approximately its criticaltemperature, which is the temperature corresponding to the criticalpoint, being the combination of pressure and temperature at which theintensive properties of the gas and liquid phases of the matter areequal, and approximately its cricondentherm, which is the maximumtemperature above which the natural gas cannot be formed into the liquidphase, regardless of pressure. In the dense phase, a gas has a viscositysimilar to that of the gas phase, but can have a density closer to thatof the liquid phase.

As used herein, the term “non-condensable” means any gas that does notliquefy at the operating pressure and temperature of a specific stage orstages for any LPS within the flow path.

As used herein, the term “absorption refrigeration process” or “ARP”,means a refrigeration system that utilizes the art recognizedthermodynamic refrigeration process that is based on thermal input todrive a refrigeration process.

As used herein, the term “trim heat” means heat input into a systemoriginating from any means of waste heat recovery, heat transfer medium,electrical resistance heaters, or other conventional means of providingheat input to a modified ARP rich solution heating loop of the presentinvention. Trim heat is preferably supplied from low-grade heat sources.Low-grade heat means low- and mid-temperature heat that has less energydensity and cannot be converted efficiently by conversional method.Although there is no unified specification on the temperature range oflow-grade heat, it is understood that a heat source with temperaturebelow 370 C is considered as a low-grade heat source, because heat isconsidered not converted efficiently below that temperature using steamRankine cycle. The main low-grade heat sources are from solar thermal,geothermal, and industrial waste thermal,

As used herein, the term “mechanical refrigeration process”, means arefrigeration system that utilizes the art recognized thermodynamicrefrigeration process that is based on compression input to drive therefrigeration process.

As used herein, the term “turbo-expander refrigeration process”, means arefrigeration system that utilizes the art recognized thermodynamicrefrigeration process that is based on adiabatic expansion and recoveryof work for compression as a refrigeration process.

In one aspect, embodiments of the present invention comprise a systemwhich comprises a gas receiving stage, a chilling stage, a liquefactionstage or stages, and a modified ARP which drives the chilling stage. Ina preferred embodiment, the invention may also comprise a compressionstage, a heat of compression energy recovery stage or stages, and a gasrecycle stage. One embodiment of the present invention seeks to utilizethe potential energy (enthalpy) of an inlet gas stream and to recoverheat of compression energy during a compression stage of theliquefaction process to improve the overall thermodynamic efficiency ofa gas liquefaction process.

In one embodiment, as shown schematically in FIG. 1, the inventioncomprises a gas liquefaction system which is combined with a modifiedaqua ammonia absorption refrigeration system. The heat of compressionenergy generated as a result of compression work on the gas or gasmixture to be liquefied may be recovered by utilizing aqua-ammonia toabsorb heat from the working fluid gas stream by means of a heatexchanger. Conventional gas processing techniques reject this highquantity, low grade heat energy to the environment either through airfin fan or water cooling systems. Embodiments of the present inventionutilize the recovered heat of compression energy in an absorptionrefrigeration cycle that provides refrigeration cooling to permitliquefaction of gases.

The potential energy (enthalpy) available in the gas to be liquefied isdirectly related to the pressure and temperature of the gas as it entersthe system and is utilized during pressure reduction refrigerationprocesses such as the Joule-Thomson (JT) pressure reduction process tochill the gas or gas mixture by auto-refrigeration from adiabaticpressure reduction. The JT process is robust and simple and is suitablefor refrigeration with no practical limitations on operating within thegas-liquid phase envelope, and does not require the use of specialtycryogenic rotating equipment which are complex, expensive, and havepractical limitations requiring operation outside of the gas-liquidphase envelope.

With a source of heat, absorption refrigeration systems typicallyutilize less than 5% net electrical energy compared to the chillingenergy produced by the absorption refrigeration system. Low grade heatof compression energy that is recovered from compression work impartedon the gas stream being liquefied can provide some, all, or excessrefrigeration duty depending on the specific gas liquefactionapplication and the method employed for liquefaction. In applicationswhere insufficient heat energy is available to be recovered from theliquefaction cycle, additional trim heat energy in the form of otheravailable low grade waste heat streams and/or other conventional meansof heat input may be required to provide the required heat energy topermit the required refrigeration duty to be developed by the absorberrefrigeration system.

The absorption refrigeration system comprises a rectifier which usesheat energy to liberate ammonia from a rich aqua ammonia solution, and avapour absorber tower (VAT), which in one embodiment, permits a chillerto operate as low as −71° C. at a 10 kPa operating pressure. The VATdesign employs thermodynamic principals to eliminate the need forconventional mechanical vacuum pumps to achieve the desired vacuumoperating pressures. The VAT design also permits at least some, andpossibly all, recovery of the heat of solution and heat of condensationenergy as anhydrous vapour ammonia absorbs into the lean aqua-ammoniasolution at the top of VAT, and optionally, at additional entry pointsto the VAT. The solution strength and temperature increases from top tobottom in the VAT, with hydraulic head maintaining the aqua-ammoniasolution in a subcooled state until the final rich solution strength isreached. The heat of solution and condensation are maintained as usefulenergy within the rich solution, unlike conventional absorbers whichreject this energy to a heat sink.

In the receiving stage, the inlet gas stream is compressed ordecompressed to a desired pressure, which may be above or below thecritical pressure of the gas prior to starting the chilling/liquefactionprocess. If the inlet gas stream is above the desired pressure, it maybe throttled with a JT valve to initiate the process at a lowertemperature. In such cases, no heat of compression is recovered totransfer to the modified ARP.

In one embodiment, a method is adapted for liquefying a gas which has aninlet pressure below the critical point for the gas. The method utilizesa compressor (one or more stages), a heat of compression energy thermalrecovery system, a modified ARP, one or more JT valves, one or more LPSvessels, and a recycle gas refrigeration compressor with one or morestages. This method of liquefaction of a gas reduces or eliminates theneed for a second set of refrigeration compressors utilizing aconventional mechanical refrigeration system such as the Carnot cycle.The gas being liquefied acts as a heat transfer fluid as the vapourphase component as a result of the JT flash is recycled and the liquidphase component sent to storage. This example of the method may besuitable for liquefaction of CO₂, H₂S, propane, or shallow cut C₃+natural gas liquid (NGL) recovery, where the required temperature forliquefaction is warmer than −70° C.

For example, this method may be used to liquefy CO₂ gas, shownschematically in FIG. 2 as a PFD, and in FIG. 3 as a Mollier Chart.Typical liquid CO₂ storage range is between about −15° C. and −29° C.The process may produce liquid CO₂ at a temperature of about −23° C., ata pressure of about 1600 kPa. At the inlet, CO₂ is delivered at anatmospheric pressure and at about 30° C., well below the critical pointof the gas. The gas is then compressed in stages while passing throughheat exchangers which recover the heat of compression energy with heatexchangers in direct communication with a rich aqua-ammonia solution, toprovide all or a portion of the heat energy required to power themodified absorption refrigeration chiller system. The compressed CO₂ isthen chilled by at least one absorption chiller. The heat energy forpowering the absorption chiller system is provided by any combination ofrecovered heat of compression energy and/or trim heat, which may beproduced by direct or indirect combustion heat exchange, or otheravailable waste heat recovery streams with the necessary temperature andmass flow conditions.

The compressed and chilled CO₂ is then released through a JT valve intoa low pressure separator (LPS) at a release pressure and a releasetemperature such that the CO₂ is in a two phase gas-liquid state, whichmay under some circumstances be in a sub-cooled state. Liquid CO₂ can bedischarged to a storage vessel, while the gas portion comprising anyflash gases and/or non-condensable vapours is directed to the recyclecompressor, a bleed stream for venting, fuel gas and/or additionalprocessing as the case may be. The recycle compressor is part of arecycle loop where the gas portion is introduced into the gas flowpathat the compression stage, as is seen in FIG. 2 at CO2-11 and CO2-11 a.

In another embodiment, the present invention provides a method forliquefying a gas which is received above its critical point, or iscompressed to above its critical point utilizing a compressor (one ormore stages), a heat of compression energy thermal recovery system, amodified ARP, one or more JT valves, one or more LPS vessels, and arecycle gas refrigeration compressor with one or more stages. The methodutilizes a flow path including but not limited to a refrigeration cycleutilizing compression of a gas to a pressure sufficiently into the densephase to permit liquefaction by means of cooling the dense phase gaswith any combination of a heat of compression energy thermal recoverysystem, an absorption refrigeration system, and heat exchanger with thelow pressure recycle gas vapour stream from one or more LPSs, one ormore JT valves, and recycle gas refrigeration compressor with one ormore stages of compression. Depending on the feed pressure andtemperature of the gas entering the liquefaction process it may beadvantageous to compress the gas further into the dense phasesufficiently above the critical pressure and temperature of the gas tooptimize the heat removed (enthalpy change) during the chilling processfor the specific gas or gas mixture to be liquefied.

The pressure selected for the chilling process for a specific gas or gasmixture is directly related to the slope change of the isotherm for thegas or gas mixture above the critical point as presented on a Pressureversus Specific Enthalpy Mollier Chart, The point at which the slope ofthe isotherm is vertical (infinite slope) provides the maximum potentialfor sensible heat transfer to occur for a given gas or gas mixture at agiven temperature. The actual pressure selected may not necessarily bethis point as a combination of factors are necessary to be consideredsuch as practical pressure and temperature limits for compression andheat exchange equipment and the minimum temperature available orprovided by the absorption chilling system. The slope of the isothermfor a specific gas can be observed on a Mollier Chart (X-axis SpecificEnthalpy and Y-axis Absolute Pressure) to assist in selection of theoptimum pressure for the chilling of a given gas or gas mixture prior tothe liquefaction step. This selection process will be described furtherbelow.

In examples of this embodiment, the gas is received at an inlet pressureat a desired dense phase pressure, or if the inlet pressure is not atthe desired dense phase pressure, compressing or decompressing the gasto the desired dense phase pressure required for liquefaction. Ifcompressed, the heat of compression energy may be recovered by means ofa heat exchanger and transferred to a rich aqua-ammonia solution, toprovide all or a portion of the heat energy required to power themodified absorption refrigeration chiller system. If the heat energyrecovered from the heat of compression is insufficient, trim heat may beprovided by any direct or indirect combustion heat exchange, or otheravailable waste heat recovery streams with the necessary temperature andmass flow conditions,

The inlet gas may be compressed in a single or multi-stage compressor asrequired to reach the desired final pressure, equal to the inletpressure of the JT Valve. Generally, in one embodiment, the dischargetemperatures for any particular compression stage is limited to about150 to 160° C., depending on the specific compression equipmentspecifications.

The compressed gas is chilled by means of at least one, and preferably2, 3 or 4 stages, absorption chiller to a minimum temperature of −70° C.In one embodiment, the compressed gas may be initially chilled with alow pressure vapour recycle stream from the LPS as discussed below.

The chilled industrial gas or gas mixture is released through a JT valveinto a low pressure separator (LPS) at a release pressure and a releasetemperature such that the gas is at a vapour quality “X” within the twophase region for the gas. In FIGS. 6A and 6B, X=0.53 at M17 and M12respectively. The liquid is discharged from the LPS to a storage vesseland the vapour is directed to a low pressure vapour recycle stream. Thisrecycle stream incorporates a heat exchanger that initially cools thedense phase industrial gas or gas mixture to the desired temperatureprior to chilling the dense phase gas in the absorption chiller. The lowpressure vapour recycle stream is thereby warmed to a temperaturesuitable for inlet to the compression equipment, and is then compressedin one or more stages until the desired dense phase liquefactionpressure is reached and then combined with the inlet gas stream.

Non-condensable vapours from the LPS may be directed to a bleed streamfor venting, additional processing or as a fuel gas depending on theproperties of the specific industrial gas or gas mixture and processapplication.

Additional stages for flashing of the liquid removed from the LPS may beconducted to further reduce the temperature and pressure of theliquefied industrial gas or gas mixture if desired with the use ofadditional JT valves, LPS vessels, and compression stages as desired.

For some gases, the absorption refrigeration chillers do not operate ata sufficiently low temperature to permit simple JT flashing of thechilled dense phase fluid to a sub-cooled state at the desired finalliquefaction temperature, but does permit flashing to the desired finaltemperature and pressure to a certain vapour quality “X” within thegas-liquid phase envelope for the gas. The liquid portion is removedfrom the LPS and sent to a liquid storage vessel, while the gas phase isremoved from the separator and the cold low pressure gas phase may beused to further cool the warmer dense phase gas stream which has beenchilled in the final stage absorption chiller heat exchanger.

In one embodiment, by cooling the dense phase gas stream, the lowpressure vapour recycle stream from the LPS is warmed to a temperatureapproaching the final absorption chiller operating temperature. It maythen be directed to another heat exchanger which further warms therecycle gas in a compressor loop to a temperature acceptable for theselected recycle compressor equipment (−29° C. or warmer to utilizestandard nodular iron or carbon steel materials and avoid the need forstainless steels necessary for cryogenic operations). Once the lowpressure vapour recycle gas has exchanged sufficient energy and issuitably warmed it may be combined with the inlet gas stream andcompressed as described herein.

Depending on the specific application, there may be excess chillingavailable for other processes, or there may be additional trim heatenergy required to permit sufficient chilling duty to be generated bythe absorption refrigeration chilling equipment. After the heat ofcompression energy has been recovered from the inlet stream and recyclestream (combined flow is equal to the inlet flow+recycle flow (“Y”)),this stream is cooled further by one or more stages of the absorptionchilling system until the desired final temperature from the absorptionchilling system is reached. The low pressure vapour recycle stream fromthe LPS has a mass flow equal to “Y” or “X/(1−X)” times the inlet flowto be liquefied. The liquid mass flow leaving the LPS is equal to theinlet mass flow of the gas or gas mixture entering the system, less anyfuel gas or bleed stream to prevent build-up of non-condensable gases atthe desired system liquefaction conditions.

Methods of gas liquefaction described herein may minimize the need foradditional equipment that are required by conventional refrigerationprocesses with cascaded multi-stage external refrigeration processes ormixed refrigerant systems that are currently utilized in large scale LNGliquefaction facilities and which require significant net energy inputand capital to construct and working capital to operate and maintain.Additionally, brazed aluminium heat exchanger (BAHX) and cryogenicrotating equipment are not required.

Additional JT flash stages may be added if colder and lower pressureliquefied gas or mixed gas products are desired which result inadditional recycle or gas bleed steams. Depending on the properties ofthe gas or gas mixture being liquefied, it may be desired to use one ora combination of the vapour streams for fuel gas or as a feed stream forrecovery of the non-condensable gases in another liquefaction process atalternate operating pressure and temperatures that permit liquefactionof the non-condensable gas or gas mixtures. One example of this methodis suitable for applications with liquefaction temperatures as low as−170° C. and is particularly suitable for LNG production or deep cut C₂+recovery.

Methods described above which use a dense phase gas are capable ofcooling a gas to a temperature of −71° C. prior to adiabatic expansion,which is sufficient to liquefy methane. In another embodiment, where therequired liquefaction temperature is lower, the invention may comprisean additional cooling step, where the vapourization of a separateliquefied industrial gas further cools the gas desired to be liquefied.This method for liquefying gases utilizes a compressor (one or morestages), a heat of compression energy thermal recovery system, amodified ARP, one or more JT valves, one or more LPS vessels, arefrigeration recycle compressor with one or more stages, and one ormore liquefied gas vapourizer heat exchangers.

In this embodiment, a liquefied gas is produced using the stepsdescribed above and further adds the step of utilizing an liquefied gasvapourizer heat exchanger to chill another dense phase gas from thefinal stage modified absorption chiller temperature to a sufficientlylow temperature that the chilled dense phase gas can further be chilledwith the recycle vapour stream from the LPS to permit liquefaction ofthe industrial gas or gas mixture by JT adiabatic expansion to a vapourquality “X” at the desired temperature and pressure. If LNG is used inthe liquefied gas vapourizer, then approximately one kg of air may beliquefied for every 0.35 kg of LNG vapourized using an embodiment of themethod illustrated in FIGS. 9 and 10.

Accordingly, stages for alternative embodiments are similar but maydiffer in required operating temperature, pressures, and heat andmaterial balances for the gas liquefaction applications. Solutionconcentration of the lean and rich aqua-ammonia solution concentrationsand flow rates is dependant primarily on the ambient (heat sinktemperature) and desired final chiller stage operating temperature.Circulation rate of a given solution mix is dependent on total coolingload required and available heat input to the system. Calculation anddetermination of these parameters are well within skill of an ordinaryskilled artisan having the benefit of this disclosure.

One feature of the present invention comprises the recovery of asignificant amount and in some cases all of the heat of solution andheat of condensation energy in the VAT, which heat is rejected to theambient environment or heat sink in conventional ARP configurations.Another feature of one embodiment of the VAT segment of the invention isthat it may achieve very low −71° C. chilling in the final chiller stagewith no requirement for rotating vacuum pump equipment, thus providing asimpler robust lower capital cost solution to achieve liquefaction ofLNG with minimal rotating equipment, and in particular no cryogenicrotating equipment.

EXAMPLES

The following examples are described to illustrated specific embodimentsof the claimed invention, and are not intended to limit the claimedinvention.

In FIGS. 2 and 3, CO₂ gas is received at atmospheric pressure and atabout 30° C., and is then compressed to a pressure of about 4400 kPa,through three stages of compression (STG-1, STG-2, and STG-3), whilebeing cooled with heat exchangers (WHX-1, WHX-2 and WHX 3). The gas isthen chilled initially by vapour recycle stream from the final separator(MP Sep) and then an absorption chiller (NH3-CH1(10)). WHX-1, WHX-2 andWHX 3 transfers heat to the aqua ammonia system, to power the absorptionchiller system

The chilled CO₂ then passes through a JT valve into the separator (MPSep) at a release pressure and a release temperature such that the CO₂is in the two phase gas-liquid, which may under some circumstances be ina sub-cooled state. The liquid portion is discharged to a storagevessel, while and the gas portion comprising any flash gases and/ornon-condensable vapours is directed to the recycle compressor, a bleedstream for venting, fuel gas and/or additional processing as the casemay be;

FIGS. 4 and 5 shows PFDs shows a liquefaction method and system forliquefying sweet natural gas, while FIG. 6 shows a Mollier Chart for thenatural gas flowpath. Table 1 attached summarizes the heat and materialbalance for these examples.

FIG. 8 depicts the flow path of the modified ARP and VAT on a PTX graphat pressures down to 10 kPa, developed in order to depict the flowpathof this example. Conventional PTX graphs for aqua-ammonia generally donot extend below 100 kPa and do not take into account operation of ARPsystems operating below atmospheric pressure. FIG. 8 depicts theflowpath of the anhydrous ammonia 4 stage gas chilling system to permitoperation at the top of the VAT at pressures as low as 10 kPa and −71°C. Table 1 summarizes the properties of the gas, methanol, aqua-ammoniasolution, and anhydrous ammonia as they progress through the flow pathobtained from available Mollier charts for methane and anhydrousammonia, publically available tables, graphs and charts for thethermodynamic properties of aqua-ammonia solutions, vapour, and steamtables. Calculations for the expected performance and operatingparameters for the modified ARP and VAT were developed by the inventoras part of the invention. The hand calculations are subject to rounding,simplification, estimating and approximation as necessary to develop thekey parameters and key system operating parameters. For example,non-condensable gases were excluded and pure anhydrous ammonia wasassumed to simplify the required hand calculations (versus target 99.5%purity). Mathematical simulation using process simulation software mayresult in refinement of calculations to permit detailed process designof required bleed streams for the gas liquefaction loop and the modifiedARP and VAT system.

A method for the liquefaction of water saturated sweet natural gas (98%CH₄ and 2% CO₂) is shown in FIGS. 4, 5 and 6A. The natural gas issupplied into the flow path at an inlet separator at M1 at a pressure of170 kPa and 17° C. The gas is compressed in the 1^(st) stage inlet(COMP-IN) and compressed to 650 kPa (M2) the same pressure as the firststage recycle gas (STG-1, M3). The heat of compression from the 1^(st)stage inlet is recovered in WHX-IN (M2 to M2 a), the heat of compressionfrom the first stage inlet is (M3 to M3 a) is used to warm the recyclegas (M20 to M21) to at least −29° C. which is the minimum acceptabletemperature for operation in a compressor of standard materials ofconstruction (non-cryogenic). The combined temperature to the inlet ofthe suction of the 2^(nd) stage recycle compressor is 47° C. (M4). Thetemperature and pressure at M4 must reviewed to ensure that hydrates orfreezing are not an issue, for this example there is not an issue butrecycle ratios to inlet gas and water content can change depending onthe application.

The combined inlet and recycle gas are compressed in the 2^(nd) recyclestage to 2,200 kPa (M5), the gas is cooled and the heat of compressionrecovered in WHX-1 (M5 to M6). The gas is further compressed in the3^(rd) stage recycle (STG-3) and heat of compression recovered startingat 160° C. in WHX-2 (M7 to M8) to 47° C.

The gas now enters a point in the flow path for pre-treatment inpreparation of the liquefaction process. HSX-5 is utilized to providecontrol of the temperature in advance of the gas in the flowpathentering the Amine Contactor (M9) where the CO₂ content is reduced from20,000 ppm to less than 50 ppm to prevent solidification of CO₂ in theliquefaction process. The gas exiting the amine contactor at point M10is water saturated as it enters the TEG glycol dehydrator, where exitingat point M11 the water vapour content has been reduced to 0.065 kg/10³m³. At point M12, 11.7 kg of methanol is injected to ensure a roughly75/25 methanol/water mixture as condensation occurs along the flowpathto the HPS at 8,200 kPa and −88° C. (M14). The condensed methanol andwater mixture is removed from the HPS (M16) leaving a dehydrated vapourstream with trace amounts of MeOH/H₂O that will solidify as the gas isflashed across the JT-1 valve to 170 kPa and −152° C., and can beremoved by filtering the liquid product. This technique for dehydrationhas not been utilized or recognized in prior art as a method ofdehydration of the gas stream during the production of cryogenic gasessuch as LNG and is a method that may be utilized in the invention toeliminate the need for conventional molecular sieve dehydration units.

Returning to M12 the point in the flowpath for methanol injection thegas is sequentially cooled from 22° C. (M12) to −88° C. (M14) viaammonia chillers NH₃-CH1, NH₃-CH2, NH₃-CH3, NH₃-CH4, and GGX-2. In thisexample, due to the warmer ambient temperature, HSX-3 does not provideany beneficial heat transfer duty. In colder climates, HSX-3 may providesignificant cooling, which will reduce the chilling duty of NH3-CH1. Asa result, thermal efficiency of the gas liquefaction process willimprove as the ambient temperature declines during winter months.

Once the chilled dense phase gas has reached the HPS, the condensedMeOH/H₂O liquid is removed at point M16 as noted previously, thedehydrated chilled high pressure gas leaves the HPS at point M15 (−88°C., 8,200 kPa) and is flashed across the Joule Thomson valve JT-1 topoint M17 (170 kPa, −152° C., and a vapour fraction X=0.53) into theLPS. The liquid LNG is removed from the LPS via M-18 by gravity to theLNG storage system (with trace MeOH/H₂O solids filtered and removed fromthe LNG), and the cold recycle gas vapour is recycled back to act as aheat transfer fluid, cooling the gas stream in GGX-2 (M13 to M-15) andwarming from M-19 to M20 (−152 to −71) the close approach temperaturesare obtained utilizing a high pressure cryogenic heat exchanger. Therecycle gas is further warmed in GGX-1, a lower pressure cryogenic heatexchanger to a minimum of −29° C. to permit the use of non-cryogeniccompression equipment, which may be either reciprocating or centrifugalas the size of the gas liquefaction plant increases.

The rich solution is received at the inlet to the rich aqua-ammoniasolution pump at point Aq1 from the bottom of the VAT in a subcooledstate of 50° C. or less and 10.4 wt % for this application. Warmerambient conditions versus cold winter ambient conditions result in loweroverall rich and lean solutions being utilized for the modifiedadsorption ARP. In this example the lean concentration is 5 wt % and therich is 10.4 wt %.

The VAT receiving anhydrous ammonia vapour from the four gas chillers(NH₃-CH1, NH₃-CH2, NH₃-CH3, NH₃-CH4) in this example operates at 10 kPaat the top and a lean solution subcooled temperature of 22° C.

Generally as colder ambient temperature or heat sinks are available thechiller duty for NH3-CH1 duty is reduced, in this example as the HSX-3cannot reduce the flowpath temperature at M12 a below 22° C., it's dutyis minimal. Lower ambient temperatures also result in the condensingoperating pressure of HSX-2 (anhydrous ammonia condenser) being reduced.The sensitivity of the reduction in ambient temperature on the modifiedARP and VAT can be seen graphically on FIG. 8 (the PTX diagram forabsorber operating pressures down to 10 kPa). Lower ammonia condensingpressure (HSX-2) and reduced duty load on the lean aqua-ammonia solution(NH3-CH1) provide opportunities to further optimize rich and leansolution strengths and circulation rates.

Returning to the rich aqua-ammonia solution pump, the discharge pressureof the pump is a direct function of the condensing temperature (andpressure) of the ammonia condenser (HSX-2). In this example 950 kPapressure is required for Aq-2, the 10.4 wt % rich solution at this pointin the flow path is subcooled. The rich solution flows first to the heatof compression recovery step splitting in parallel with flow rates splitproportionate to the waste heat recovery duty of each exchanger (WHX-IN,WHX-1, WHX-2) rising in temperature from 50° C. (Aq2) to 72.5° C. (Aq10)at 10.4 wt % and 940 kPa the rich solution is still subcooled.

The next point in the flow path is the rich/lean solution exchangerwhere the rich solution is further heated to 143° C. at which point Aq12the rich solution enters the modified ARP rectifier column.

As a result of the operating pressures at the 22° C. condensingtemperature (HSX-2), the modified ARP system for this example iscalculated to have a trim heat requirement of 924 kW, which can besupplemented from available low grade waste heat recovery streams, butrequires an ultimate final temperature of 159° C. to achieve the leansolution concentration of 5 wt %. The additional waste heat could besupplied directly to the generator/surge vessel or along the richsolution heat exchanger heating loop.

Once sufficient additional trim heat is provided the required reflux andvapour traffic will be achieved in the rectifier column. Thedephlegmator DPX requires 436 kW of cooling duty to achieve a 50° C.exit temperature which results in an ammonia stream that is anticipatedto be 99.5 wt % ammonia based on the assumed reflux ratio of 2 and alean saturated solution strength of 5 wt % (Aq14).

The lean solution Aq14 is subcooled in the lean rich solution exchangerand the temperature is reduced from 159° C. to 85° C. (Aq15), The leansolution is further cooled in HSX-1 to 22° C. in this example at pointAq16 in the flow path. The subcooled 5 wt % lean solution is injectedinto the top of the VAT column approximately 10.6 m elevation higherthan the Rich Aq pump suction, The lean aqua-ammonia solution at 5 wt %is subcooled at 22° C. to permit the ammonia from NH3-19 in the flowpath (at −71° C. and 10 kPa) to fully dissolve in the subcooled leansolution and to remain in a subcooled state after accounting for therise in temperature from heat of solution and heat of condensationenergy and enthalpy mixing of the ammonia vapour and lean solution.

The 10 kPa operating pressure is developed by pinching the leanaqua-ammonia flash valve thereby reducing the pump suction pressure ofthe rich aqua-ammonia pump but maintaining the suction pressure abovethe NPSHR and a subcooled lean solution to ensure absorption of theanhydrous ammonia vapours,

Returning back to the point in the flow path where the ammonia vapour(NH3-1) exits the top of the DPX, the vapour continues to the ammoniacondenser (HSX-2). It is the condensing temperature of this heatexchanger that sets the operating pressure for the rich solution side ofthe modified ARP. The HSX-2 removes 230 kW to condense the requiredammonia vapour flow for this example,

After NH3-2, the ammonia is fully condensed and FIG. 7 the anhydrousammonia Mollier Diagram details the thermodynamic aspects of thisportion of the invention, The ammonia is at 900 kPa for the 22° C.condensing pressure and is flashed to the corresponding pressures forthe 4 chillers as shown in the PFD for the modified ARP and VAT aspectsof the invention, Shown on FIG. 5 associated with the ammonia chillersare NH₃ Bleed Valves 1, 2, 3 and 4, For the purposes of simplified handcalculations in this example, bleed streams are not utilized. However,in practice a bleed stream of approximately 5% may be required for eachammonia chiller to prevent a build-up of H₂O in the ammonia chillers,which may render the system non-functional, The actual bleed stream willdepend on the purity of the ammonia produced from the rectifier column,which for this example was targeted for 99.5% purity.

At the top of the VAT, the ammonia entering the VAT is at a height,temperature, mass flow rate that results in the aqua-ammonia solutionincreasing in strength and temperature as the solution flows down theVAT. As in shown on FIG. 8, the PTX chart the solution remains subcooledin this example for flows NH3-19 (Aq17 24.6° C., 10 kPa, 5.5 wt %),NH3-15 (Aq18 28.1° C., 13 kPa, 6.2 wt %), NH3-11 (Aq19 34.1° C., 30 kPa,7.3 wt %), and NH3-6 (Aq20, 49.8° C., 72 kPa, 10.4 wt %) for the fourgas flowpath chillers. If a superheated solution were to occur at thelowest mixing point (Aq20), a heat exchanger HSX-4 could be employed toremove excess heat to subcool the rich aqua-ammonia solution prior topump suction (Aq1) to maintain the desired operating pressures at thetop of the VAT.

In another example, shown in FIGS. 8 and 9 which illustrate theliquefaction of air, the inlet gas is delivered at a pressure below thecritical point. Liquefied air is produced by utilizing a liquefied gasvapourizer to provide additional cooling in the flow path downstream ofthe final stage absorber chiller (which operates at −70° C.) in order topermit a temperature and pressure condition to be reached that resultsin a flashed gas or gas mixture at the desired temperature and pressureto be within the gas-liquid phase envelope at a certain quality “X”.

For example, natural gas may be liquefied using the methods describedabove, and then the LNG could be vapourized to provide additionalchilling to the air stream beyond the chilling provided by the finalchiller stage of a modified absorption chilling system. The vapourizednatural gas may then become the feed for the LNG liquefaction looputilizing an alternative embodiment as described above, or as a sourceof gaseous fuel if the air liquefaction plant was co-located on a siteutilizing LNG as a source of fuel. This method may be suitable forliquefaction of a gas requiring very low temperatures (lower than −170°C.) to enable liquefaction to occur, and minimize additional equipmentthat is required by conventional refrigeration processes with cascadedmulti-stage external refrigeration processes.

DEFINITIONS AND INTERPRETATION

All references to temperatures and pressures in the description hereinshould be considered to be modified with the term “about”, which means avariation of ±5%, ±10%, ±20%, or ±25% of the value specified. Forexample, “about” 50 percent can in some embodiments carry a variationfrom 45 to 55 percent. For integer ranges, the term “about” can includeone or two integers greater than and/or less than a recited integer ateach end of the range. Unless indicated otherwise herein, the term“about” is intended to include values and ranges proximate to therecited range that are equivalent in terms of the functionality of thecomposition, or the embodiment described. The term “about” may alsoreflect any imprecision in instruments, devices or methods used tomeasure the value specified.

As will be apparent to those skilled in the art, various modifications,adaptations and variations of the foregoing specific disclosure can bemade without departing from the scope of the invention claimed herein.The various features and elements of the invention described herein maybe combined in a manner different than the specific examples describedor claimed herein without departing from the scope of the invention. Inother words, any element or feature may be combined with any otherelement or feature in different embodiments, unless there is an obviousor inherent incompatibility between the two, or it is specificallyexcluded.

The singular forms “a,” “an,” and “the” include plural reference unlessthe context clearly dictates otherwise. Thus, for example, a referenceto “a plant” includes a plurality of such plants. It is further notedthat the claims may be drafted to exclude any optional element. As such,this statement is intended to serve as antecedent basis for the use ofexclusive terminology, such as “solely,” “only,” and the like, inconnection with the recitation of claim elements or use of a “negative”limitation. The terms “preferably,” “preferred,” “prefer,” “optionally,”“may,” and similar terms are used to indicate that an item, condition orstep being referred to is an optional (not required) feature of theinvention.

The term “and/or” means any one of the items, any combination of theitems, or all of the items with which this term is associated. Thephrase “one or more” is readily understood by one of skill in the art,particularly when read in context of its usage.

As will be understood by the skilled artisan, all numbers, includingthose expressing quantities of reagents or ingredients, properties suchas molecular weight, reaction conditions, and so forth, areapproximations and are understood as being optionally modified in allinstances by the term “about.” These values can vary depending upon thedesired properties sought to be obtained by those skilled in the artutilizing the teachings of the descriptions herein. It is alsounderstood that such values inherently contain variability necessarilyresulting from the standard deviations found in their respective testingmeasurements.

As will be understood by one skilled in the art, for any and allpurposes, particularly in terms of providing a written description, allranges recited herein also encompass any and all possible sub-ranges andcombinations of sub-ranges thereof, as well as the individual valuesmaking up the range, particularly integer values. A recited range (e.g.,weight percents or carbon groups) includes each specific value, integer,decimal, or identity within the range. Any listed range can be easilyrecognized as sufficiently describing and enabling the same range beingbroken down into at least equal halves, thirds, quarters, fifths, ortenths. As a non-limiting example, each range discussed herein can bereadily broken down into a lower third, middle third and upper third,etc.

As will also be understood by one skilled in the art, all language suchas “up to”, “at least”, “greater than”, “less than”, “more than”, “ormore”, and the like, include the number recited and such terms refer toranges that can be subsequently broken down into sub-ranges as discussedabove. In the same manner, all ratios recited herein also include allsub-ratios falling within the broader ratio. Accordingly, specificvalues recited for radicals, substituents, and ranges, are forillustration only; they do not exclude other defined values or othervalues within defined ranges for radicals and substituents.

One skilled in the art will also readily recognize that where membersare grouped together in a common manner, such as in a Markush group, theinvention encompasses not only the entire group listed as a whole, buteach member of the group individually and all possible subgroups of themain group. Additionally, for all purposes, the invention encompassesnot only the main group, but also the main group absent one or more ofthe group members. The invention therefore envisages the explicitexclusion of any one or more of members of a recited group. Accordingly,provisos may apply to any of the disclosed categories or embodimentswhereby any one or more of the recited elements, species, orembodiments, may be excluded from such categories or embodiments, forexample, as used in an explicit negative limitation.

TABLE 1 M1 M2 M2a M3 M3a M4 M5 M6 M7 M8 M9 M10 M11 Stream Name -Liquefier - Methane Loop

Pressure (kPa) 170 650 600 650 600 600 2200 2100 8600 8520 8500 84508400 Temperature (° C.) 17 123 65 69 32 47 157 47 160 47 22 22 22Density Vapor (kg/m3) 1.02 3.3 3.25 3.9 4 3.7 10.5 12.5 38 57 61 61 61Density Liquid (kg/m3) Vapor Fraction 1 1 1 1 1 1 1 1 1 1 1 1 1 MeOH (wt% of liquid) 0 0 0 0 0 0 0 0 0 0 0 0 0 Enthalpy (kJ/kg) 892 1140 9951005 920 955 1220 940 1200 885 820 820 820 Approximate Mass in kg with28.3 e3m3/d -inlet C1 (0.98 mol vol dry) (kg/d) 18500 18500 18500 2086020860 39360 39360 39360 39360 39360 39360 39360 39360 C02 (0.02 mol voldry) kg/d) 575 575 575 0 0 575 575 575 575 575 575 0.5 0.5 H2O (vapor)(kg/d) 250 250 250 0 0 250 250 250 250 18 18 18 3.9 H2O (liquid) (kg/d)0 0 0 0 0 0 0 0 0 232 0 0 0 MeOH (kg/d) 0 0 0 0 0 0 0 0 0 0 0 0 0Estimate of TOTAL mass flow (kg/d) 19325 19325 19325 20860 20860 4018540185 40185 40185 40185 39953 39379 39364 mass flow (kg/s) 0.224 0.2240.224 0.241 0.241 0.465 0.465 0.465 0.465 0.465 0.462 0.456 0.456 M12M12a M12b M12c M12d M13 M14 M15 M16 M17 M18 M19 M20 M21 Stream Name -Liquefier - Methane Loop

Pressure (kPa) 8400 8300 8275 8200 8225 8200 8200 8200 8200 200 200 200185 170 Temperature (° C.) 22 22 −39 −54 −64 −70 −88 −88 −88 −152 −152−152 −70 −29 Density Vapor (kg/m3) 61 61 105 155 205 245 300 300 3.3 3.31.6 1.3 Density Liquid (kg/m3) 245 300 800 410 410 Vapor Fraction 1 1 11 1 0.999 0.999 1 0 0.53 0 1 1 1 MeOH (wt % of liquid) 0 0 0 0 0 0.750.75 0.75 0.75 0 0 0 0 0 Enthalpy (kJ/kg) 820 820 610 500 435 385 290285 285 30 525 705 790 Approximate Mass in kg with 28.3 e3m3/d -inlet C1(0.98 mol vol dry) (kg/d) 39360 39360 39360 39360 39360 39360 3936039360 0 39360 18500 20860 20860 20860 C02 (0.02 mol vol dry) kg/d) 0.50.5 0.5 0.5 0.5 0.5 0.5 0.5 0 0.5 0.5 0 0 0 H2O (vapor) (kg/d) 3.9 3.93.9 3.9 3.9 0.04 0.04 0.04 0 0.04 0.04 0 0 0 H2O (liquid) (kg/d) 0 0 0 00 3.86 3.86 0 3.86 0 0 0 0 0 MeOH (kg/d) 11.7 11.7 11.7 11.7 11.7 11.711.7 0.12 11.58 0.12 0.12 0 0 0 Estimate of TOTAL mass flow (kg/d)39376.1 39376 39376 39376 39376 39376 39376 39361 15.44 39361 18500.6620860 20860 20860 mass flow (kg/s) 0.456 0.456 0.456 0.456 0.456 0.4560.456 0.456 0.000 0.456 0.214 0.241 0.241 0.241 Aq1 Aq2 Aq3 Aq4 Aq5 Aq6Aq7 Aq8 Aq9 Aq10 Aq11 Aq12 Aq13 Aq14 Stream Name - Adsorption Refrig -Aq-N

Pressure (kPa) 110 950 950 950 950 950 940 940 940 940 930 900 900 910Temperature (° C.) 50.0 50.0 50.0 50.0 50.0 50.0 72.5 72.5 72.5 72.5143.0 50.0 50.0 160.0 Temperature (° F.) 122.0 122.0 122.0 122.0 122.0122.0 162.5 162.5 162.5 162.5 289.4 122.0 122.0 320.0 Subcooled yes yesyes yes yes yes yes yes yes yes Saturated yes yes yes yes Suprerheatedwt % Solution Aq-NH3 0.104 0.104 0.104 0.104 0.104 0.104 0.104 0.1040.104 0.104 0.104 0.560 0.050 Density Solution Aq-NH3 (kg/m3) 937 TotalEnthalpy of AqMixture 126 126 126 126 126 126 225 225 225 225 527 7 638(KJ/kg) (Note 1) Total Enthalpy of AqMixture 54 54 54 54 54 54 98 98 9898 227 3 275 (Btu/lb) (Note 1) Approximate Mass in kg with 28.3 e3scm3/d-inlet Mass Aq-Ammonia liquid (kg/s) 3.1715 3.1715 1.5030 1.6685 1.33500.3330 0.3330 1.3350 1.5030 3.1715 3.1715 2.9900 Mass Aq-Ammonia vapor(kg/s) 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.00000.0000 0.0000 0.0000 TOTAL Mass Aq-Ammonia (kg/s) 3.1715 3.1715 1.50301.6685 1.3350 0.3330 0.3330 1.3350 1.5030 3.1715 3.1715 2.9900 Aq15 Aq16Aq17 Aq18 Aq19 Aq20 Stream Name - Adsorption Refrig - Aq-NH3 LoopPressure (kPa) 890 880 10 13 30 72 Temperature (° C.) 85.0 22.0 24.628.1 34.1 49.8 Temperature (° F.) 185.0 71.6 76.3 82.6 93.4 121.6Subcooled yes yes yes yes yes yes Saturated wt % Solution Aq-NH3 0.0500.050 0.055 0.062 0.073 0.104 Density Solution Aq-NH3 (kg/m3) TotalEnthalpy of AqMixture 319 35 58 39 74 125 (KJ/kg) (Note 1) TotalEnthalpy of AqMixture 137 15 25 17 32 54 (Btu/lb) (Note 1) ApproximateMass in kg with 28.3 e3scm3/d -inlet Mass Aq-Ammonia liquid (kg/s)2.9900 2.9900 3.0058 3.0274 3.0663 3.1715 Mass Aq-Ammonia vapor (kg/s)0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 TOTAL Mass Aq-Ammonia (kg/s)2.9900 2.9900 3.0058 3.0274 3.0663 3.1715 NH3-1 NH3-2 NH3-3 NH3-4 NH3-5NH3-6 NH3-7 NH3-8 NH3-9 NH3-10 NH3-11 NH3-12 NH3-13 NH3-14 Stream Name -Adsorption Refrig - NH3 loop Pressure (kPa) 900 900 72 72 72 72 72 30 3030 30 30 13 13 Temperature (° C.) 50.0 22.0 −40.0 −40.0 −40.0 −40.0−40.0 −55.0 −55.0 −55.0 −55.0 −55.0 −65.0 −65.0 Temperature (° F.) 122.071.6 −40.0 −40.0 −40.0 −40.0 −40.0 −67.0 −67.0 −67.0 −67.0 −67.0 −85.0−85.0 Subcooled Saturated yes yes yes yes yes yes yes yes yes yes yesyes yes yes Superheated Pure Ammonia vapor fraction, 1 0 0.2 0.35 1 1 00.05 0.098 1 1 0 0.06 1 (99.5 wt %) Density Vapor (kg/m3) 5.7 0.61 0.610.61 0.61 0.27 0.27 0.27 0.27 0.13 0.13 Density Liquid (kg/m3) 610 685685 685 710 710 710 715 Enthaphy NH3-l (kJ/kg, 440 160 160 160 95 95 9550 from mollier chart) Enthapy NH3-v (kJ/kg, 1705 1550 1550 1550 15501525 1525 1525 1525 1510 1510 from mollier chart) Corrected EnthaphyNH3-l 92 −188 −188 −188 −253 −253 −253 −298 (kJ/kg) - 1938 CorrectedEnthapy NH3-v 1357 1202 1202 1202 1202 1177 1177 1177 1177 1162 1162(kJ/kg) - 1938 Approximate Mass in kg with 28.3 e3scm3/d -inlet MassNH3-l (kg/s) 0.1815 0.1452 0.0689 0.0000 0.0000 0.0763 0.0725 0.03510.0000 0.0000 0.0374 0.0203 0.0000 Mass NH3-v (kg/s) 0.1815 0.03630.0363 0.1052 0.1052 0.0000 0.0038 0.0038 0.0389 0.0389 0.0000 0.00130.0216 Mass NH3 TOTAL (kg/s) 0.1815 0.1815 0.1815 0.1052 0.1052 0.10520.0763 0.0763 0.0389 0.0389 0.0389 0.0374 0.0216 0.0216 NH3-15 NH3-16NH3-17 NH3-18 NH3-19 Stream Name - Adsorption Refrig - NH3 Loop Pressure(kPa) 13 13 10 10 10 Temperature (° C.) −65 −65 −71 −71 −71 SubcooledSaturated yes yes yes yes yes Superheated Pure Ammonia vapor fraction 10 0.018 1 1 Density Vapor (kg/m3) 0.13 0.1 0.1 0.1 Density Liquid(kg/m3) 715 725 Enthaphy NH3-l (kJ/kg, 50 30 from mollier chart) EnthapyNH3-v (kJ/kg, 1510 1500 1500 1500 from mollier chart) Corrected EnthaphyNH3-l −298 −318 (kJ/kg) - 1938 (Note 2) Corrected Enthapy NH3-v 11621152 1152 1152 (kJ/kg) - 1938 (Note 2) Approximate Mass in kg with 28.3e3scm3/d -inlet Mass NH3-l (kg/s) 0.0000 0.0158 0.0155 0.0000 0.0000Mass NH3-v (kg/s) 0.0216 0.0000 0.0000 0.0158 0.0158 Mass NH3 TOTAL(kg/s) 0.0216 0.0158 0.0155 0.0158 0.0158 Equipment Name RAqPump DPXAqRLX Trim-Heat WHX-1 WHX-2 TOTAL mass flow (kg/s) 3.172 2.990 0.4650.465 Enthalpy Change (kJ/kg) −319 −280 −315 Heat In/Out (kW) −436 −954924 −130 −147 Work In/Out (kW) 10 Equipment Name WHX-IN COMP-IN STG-1STG-2 STG-3 TOTAL mass flow (kg/s) 0.224 0.224 0.241 0.465 0.465Enthalpy Change (kJ/kg) −145 248 215 265 260 Heat In/Out (kW) −32 WorkIn/Out (kW) 55 52 123 121 Equipment Name HSX-1 HSX-2 HSX-3 HSX-4 HSX-5TOTAL mass flow (kg/s) 2.990 0.182 0.456 3.172 0.462 Enthalpy Change(kJ/kg) −284 −1265 0 0 −65 Heat In/Out (kW) −849 −230 0 0 −30 WorkIn/Out (kW) Equipment Name NH3—CH1 NH3—CH2 NH3—CH3 NH3—CH4 GGX-1 GGX-2TOTAL mass flow (kg/s) 0.456 0.456 0.456 0.456 0.241 0.241 EnthalpyChange (kJ/kg) −210 −110 −65 −50 85 180 Heat In/Out (kW) −96 −50 −30 −2321 43 Work In/Out (kW) (Note 1) Enthalphy for Aqua-Ammonia from 1938 -Jennings and Shannon Tables - Lehigh University Bethlehem, Pennsylvania(Note 2) (to zero (kJ/kg) anhydrous ammonia mollier chart to aq-Ammonia1938 tabl

 −348

indicates data missing or illegible when filed

1. A method for liquefying a gas comprising the following non-sequentialsteps: (a) receiving a gas having an inlet pressure and compressing ordecompressing the gas to a desired pressure; (b) Chilling the gasthrough at least one absorption chiller; (c) Adiabatically reducing thepressure of the gas to liquefy at least a portion of the gas; (d)heating a rich aqua-ammonia fluid in a rectifier to liberate ammonia gasusing one or a combination of trim heat or heat of compression recoveredfrom step (a) if the gas is compressed in step (a), producing a leanaqua-ammonia fluid; (e) subcooling the lean aqua-ammonia and circulatingto the top of a vapour absorption tower (VAT); (f) condensing ammoniagas from the rectifier and flashing the liquid ammonia to producechilled ammonia gas for use in the at least one absorption chiller; (g)absorbing ammonia gas from the at least one absorption chiller into thelean aqua-ammonia in the vapour absorption tower to produce the richaqua-ammonia for step (d).
 2. The method of claim 1 wherein the gas isreceived at or above the desired pressure, and no heat of compression isrecovered in step (a).
 3. The method of claim 1 wherein the gas ispartially liquefied in step (c) and comprising the further steps ofremoving the liquefied gas product and recycling the remaining gas in avapour recycle loop which cools the gas stream before adiabatic pressurereduction and is then compressed and combined with the inlet gas stream.4. The method of claim 1 wherein the gas comprises an industrial gas ora hydrocarbon gas, or any mixture of industrial or hydrocarbon gases. 5.The method of claim 1 wherein the gas is compressed below the criticalpoint of the gas and the method reaches a gas liquefaction temperatureof warmer than −71° C., prior to adiabatic expansion.
 6. The method ofclaim 1 wherein the gas is compressed to above the critical point of thegas, and the method reaches a gas liquefaction temperature of about −71°C., prior to adiabatic expansion.
 7. The method of claim 1 wherein theat least one absorption chiller comprises a liquid bleed stream toprevent increase of water concentration in the ammonia refrigerant. 8.The method of claim 1 wherein a sub atmospheric operating pressure atthe top of the VAT is maintained by utilizing a sufficient mass flow ofsubcooled lean aqua-ammonia solution at the mixing point of anhydrousammonia vapour and lean aqua-ammonia.
 9. The method of claim 1 whereinsome or all of the heat of solution and heat of condensation energy ofmixing anhydrous ammonia vapour and lean aqua-ammonia solution isrecovered in the VAT
 10. The method of claim 9 wherein heat recovery inthe VAT is facilitated by use of hydraulic head and pre-cooling of thelean aqua-ammonia VAT feed stream.
 11. The method of claim 1 comprisingthe further steps of cooling the chilled gas stream from step (b)utilizing gas/gas heat exchangers to further reduce the temperature ofgas stream prior to step (c).
 12. The method of claim 1, comprising thefurther step of dehydrating the gas after compression and before theabsorption chiller.
 13. The method of claim 12 wherein the gas isdehydrated by the addition of an alcohol in a sufficient quantity intothe flow path, and condensing the alcohol and water before the adiabaticpressure reduction step.
 14. The method of claim 1 wherein the liquefiedgas is filtered after adiabatic pressure reduction to remove any solidsfrom the liquefied gas product.
 15. A method for liquefying a gascomprising the following non-sequential steps: (a) Compressing a gashaving an inlet pressure to a desired pressure while extracting heatfrom the gas to maintain a desired temperature; (b) Chilling the gasthrough at least one absorption chiller; (c) Adiabatically reducing thepressure of the gas to reduce the temperature of the gas; (d) Liquefyingthe gas by using a liquefied gas vapourizer heat exchanger; (e) usingheat from step (a) in a rectifier to liberate ammonia from a richaqua-ammonia fluid and if necessary, adding trim heat to the rich aquaammonia, and produce a lean aqua-ammonia fluid; (f) subcooling the leanaqua-ammonia and circulating to the top of a vapour absorption tower;(g) condensing ammonia gas from the rectifier and flashing the liquidammonia to produce chilled ammonia gas for use in the at least oneabsorption chiller; (h) absorbing ammonia gas from the at least oneabsorption chiller into the lean aqua-ammonia in the vapour absorptiontower to produce the rich aqua-ammonia for step (d).
 16. The method ofclaim 15 wherein the liquefied gas vapourizer utilizes a liquefied gasproduced in accordance with the method of claim
 1. 17. A gasliquefaction system comprising a receiving stage for receiving an inletgas at a desired pressure, a chilling stage comprising an absorptionrefrigeration loop for chilling the gas, and a liquefaction stagecomprising a JT valve for at least partially liquefying the gas.
 18. Thesystem of claim 17 further comprising a compression stage forcompressing the gas to the desired pressure, a heat of compressionenergy recovery stage for transferring heat from the compression stageto the absorption refrigeration loop.
 19. The system of claim 18 furthercomprising a gas recycle stage for recycling non-liquefied components ofthe gas in a low pressure vapour recycle loop, which loop further chillsthe compressed and chilled gas, and which is then directed to thecompression stage.